Co2 and h2s removal



United States Patent 3,435,590 C0 AND H 8 REMOVAL Calvin S. Smith, ElCerrito, Calif., assignor to Chevron Research Company, San Francisco,Calif., a corporation of Delaware Filed Sept. 1, 1967, Ser. No. 665,108Int. Cl. B01d 49/00 US. Cl. 55-43 13 Claims ABSTRACT OF THE DISCLOSURE Aprocess for removing H and/ or CO from a H containing gas mixture fed toa low point in a C0 absorber. Cold, lean, liquid absorbent such aspropylene carbonate, acetone or methanol is fed to an intermediate pointin the absorber and flows downward in countercurrent contact with thegas mixture, thus absorbing H S and/or CO and cooling the H Warm, lean,liquid absorbent is fed to an upper point in the absorber and flowsdownward in countercurrent contact with the cold H at a point above theintermediate point, thus cooling the absorbent and warming the cold HThe absorbed gases are later removed by flashing and stripping.

BACKGROUND OF THE INVENTION Field of the invention The present inventionrelates to the selective removal of constituents, for example, carbondioxide and hydrogen sulfide, from gaseous mixtures containing the sametogether with other gaseous components.

Prior art A number of processes are now in use for effecting suchremoval, for example, washing (i.e., removing selected constituents byabsorption) with water under pressure or with an alkanolamine or acaustic alkali solution. It is also known to separate carbon dioxide andhydrogen sulfide from crude gases by washing and simultaneousrectification at low temperatures, preferably be tween 0 F. and -1O()F., in several stages where the wash liquid can also be regenerated byrectification.

The disadvantage of the known processes is that they are expensive. Highcapital investment for the initial plant and/or high operating costs areincurred. Recently, low temperature acid gas absorption processes usingphysical-absorption type absorbents have been put into commercialoperation, achieving a reduction in operating costs compared totraditional acid gas absorption using chemisorption-type absorbents,such as monoethanolamine. But the capital investment for the lowtemperature processes is comparatively high.

The capital investment is particularly high when it is desired toseparate gases, such as CO from H at high pressures using lowtemperature gas absorption processes. The high capital cost is largelydue to expensive high pressure heat exchangers used to cool the feed gasto the absorption column. According to present processes, typically thewarm feed gas is cooled by heat exchange against cold purified gasleaving the top of the low temperature absorber and/or cold, lowpressure CO which has been selectively separated from the H Because theheat exchange is thus generally gas to gas, relatively low heat transfercoefficients are obtained. This results in a requirement for heatexchangers with large heat transfer areas. These large heat exchangersare expensive at moderate pressures and very expensive at highpressures.

SUMMARY OF THE INVENTION According to the present invention, in aprocess for removing CO from a CO -rich gas stream containing componentshaving boiling points lower than CO where, in the CO -rich gas stream isfed to a lower point in an absorber and CO is absorbed by a downwardflowing liquid absorbent at low temperatures and high pressures, theimprovement is made which comprises:

(1) Feeding cold, lean, liquid absorbent to an intermediate point in theabsorber;

(2) Feeding warm, lean, liquid absorbent to an upper point in theabsorber; and

(3) Directly contacting cold, CO -Iean gas, flowing upward in theabsorber past the intermediate point, with warm, lean, liquid absorbentflowing downward in the absorber so as to remove heat from the absorbentand warm the cold, CO -lean gas.

In the present process, because the cold, CO -Iean gas is directly heatexchanged with warm, lean, liquid absorbent, the heat is transferred tothe cold, CO -lean gas much more efiiciently than in prior processes.The heat transfer coeflicient is much improved not only because thecold, CO -lean gas is heat exchanged with liquid (as opposed to gas-gasheat exchange), but also because the heat exchange is effected by directcontact. The direct contact eliminates the decrease in the heat transfercoefficient due to the metal wall of the tube which is between the cold,lean gas and the warm fluid when using shell-and-tube heat exchangersaccording to prior processes.

Furthermore, the CO -rich gas stream is fed to the bottom of theabsorber, warm, i.e., between 70 and 130 F., typically F., and is cooledby direct heat exchange with the downward passing absorbent while at thesame time CO is condensed and also absorbed by the absorbent. Thus thelarge high pressure exchangers used according to prior processes to coolthe feed gas are eliminated according to the present invention.

Due to the cold absorbent entering at an intermediate point, a coldpinch point is obtained in the CO -absorber and CO is forced to leavethe bottom of the absorber. The CO; laden absorbent mixture is wtihdrawnfrom the bottom of the absorber at relatively high temperatures,generally no more than 5 to 35 F. below the temperature of the feed gas.After a first reduction of pressure on the absorbent-CO mixturewithdrawn from the bottom of the absorber, heat is removed from theabsorbent-CO mixture at pressures substantially lower than the pressureof the feed gas, using cold low pressure CO obtained by subsequentfurther reduction of the pressure on the absorbent-CO mixture. Thus thegas to gas high presure heat exchange of feed gas against cold lowpressure CO as according to typical prior art processes, is avoided.

In the present invention, the absorbent is desirably a polar absorbent,and the absorption is effected at high pressure taking advantage of theincreased solubility of CO in the polar absorbent with increasedpressure. In the process of the present invention, particularly when His contained in the feed gas and the H is to be used in a high pressurehydroconversion process, it has been found economical to compress the CO-rich feed gas to above 900 p.s.i.g., usually above 1200 p.s.i.a., usingcentrifugal compressors. Using polar absorbents such as methanol,acetone or propylene carbonate at these high 'pressures results in avery efficient process for removing CO from more volatile gases, such asH Thus, according to a preferred embodiment of the invention, there isprovided a process for removing CO from H which comprises:

(1) Absorbing the CO in an absorber with cold methanol so as to obtain acold, CO -lean, H -rich gas stream and warm, CO -rich methanol;

(2) Reducing the pressure on the warm, CO -rich methanol to obtain acold, gaseous CO -rich stream and cold, CO -lean liquid methanol;

(3) Heating and stripping further CO from the cold, CO -lean liquidmethanol to obtain warm, CO lean liquid methanol;

(4) Feeding the warm, Co -lean liquid methanol to an upper section ofthe absorber; and

(5) Directly contacting the cold, CO -lean, H -rich gas stream risingupward in the absorber with the warm, CO -lean liquid methanol flowingdownward in the absorber.

The process of the present invention may be applied advantageously topuurification of raw hydrogen gas streams produced by reforming lighthydrocarbons with H O. The raw hydrogen produced by reforming followedby CO shift conversion characteristically is about 80 percent H 18percent CO 0.5 percent CO, and the balance light hydrocarbons, N Argon,etc. When the raw hydrogen is produced by reforming light hydrocarbons,centrifugal compression of the raw hydrogen to high pressures asreferred to above is particularly advantageous, as the raw hydrogen isgenerally obtained from the reforming step at pressures below 900p.s.i.a., generally about l300 p.s.i.a. according to most currently usedcommercial processes.

The gas separation process of the present invention may also beadvantageously applied to purification of raw hydrogen gas streamsproduced by partial oxidation of a wide range of hydrocarbons, such ascoal, petroleum residual fractions, heavy hydrocarbons, and lighthydrocarbons. The raw hydrogen produced by partial oxidation followed byCO shift conversion characteristically is about 60 percent H 25 percentC0 2 percent CO, 2 percent H 8 and the balance light hydrocarbons Nargon, etc. It is usually desirable to remove the H S as well as the COfrom the raw hydrogen produced. Thus according to a preferred embodimentof the invention, there is provided a process for removing CO and H Sfrom an H -containing feed gas which comprises:

(1) Absorbing the CO and H in an absorber with cold, CO H S-leanmethanol obtained as described in Step 2, so as to obtain a cold, CO HS-lean, H -rich gas stream and warm, CO H S-rich methanol;

(2) Reducing the pressure on the warm, CO H S-rich methanol so as toobtain a cold, gaseous CO and H 5 stream, and the cold, CO H S-leanmethanol used to absorb CO and H S;

(3) Scrubbing H S from the cold, gaseous CO and H 8 stream using leanmethanol obtained as described in Step 4, so as to obtain a CO -rich,gaseous stream and H S-rich methanol;

(4) Rectifying the H S-rich methanol to obtain the lean methanol used toscrub H 8 and a gaseous H S-rich stream and a warm, CO H S-lean methanolstream;

(5) Directly contacting the cold, CO H S-lean, H rich gas stream risingupward in the absorber with the warm, CO H S-lean methanol streamflowing downward in the absorber.

BRIEF DESCRIPTION OF THE DRAWINGS FIGURE 1 is a schematic illustrationof the process of the present invention as applied to a feed gascomprising hydrogen and CO and FIGURE 2 is a schematic illustration ofthe process of the present invention as applied to a feed gas comprisinghydrogen, CO and H 8.

DETAILED DESCRIPTION Referring now in more detail to FIGURE 1, feed gascomprising hydrogen and CO obtained, for example, from a steam-methanereforming process followed by shift conversion of CO to hydrogen and COand then followed by centrifugal compression from about 250 p.s.i.a. toabout 1,500 p.s.i.a. and cooling to 100 F. by a water cooled heatexchanger, is fed in line 1 to a lower point in absorber 2. At this 100F. hydrogen plus CO feed gas fed to the bottom of absorber 2 risesupward in the absorption section of the absorber, cold absorbent in line3 absorbs CO from the feed gas and cools the feed gas to within about 5to 30 F. of the cold absorbent entrance temperature at an intermediatepoint of the absorber. The cold absorbent fed to the absorber in line 3,for purposes of examplemethanol, is at a temperature between minus 30and minus 110 F., typically minus F.

Cold, CO -lean gas passes upward from the absorption section and intothe heat transfer section of the absorber. This cold, Co -lean gas isdirectly contacted with warm, CO free methanol entering the column at anupper point of the absorber as indicated by line 4, in FIGURE 1, at atemperature between 50 and 110 F., typically about 80 F. Thus the cold,CO -lean gas rising upward from the absorption section is warmed by thewarm methanol entering the absorber at the top of the heat transfersection.

Although the center portion of the column is referred to as a heattransfer section, there is also CO absorption occurring in the heattransfer section, particularly as the CO -Iean solvent is cooled. Alsoin the absorption section of the absorber there is, of course, aconsiderable amount of direct contact heat transfer between the coldabsorbent entering in line 3 and the warm feed gas entering in line 1.

Because a small fraction of the warm absorbent vaporizes and is carriedupward in the column by the CO lean gas passing upward from the heattransfer section, a water wash section is provided in the uppermost partof the absorber. Water in line 5 washes downward in the water washsection of the absorption column, thus washing down vaporized andentrained absorbent leaving the heat transfer section of the absorber.The water also serves to remove still further amounts of CO from thefeed gas because of the very low concentration of CO in the water usedfor the water wash section.

In the Water wash section, as well as the heat transfer section and theabsorption section of the absorber, bubble cap trays are one preferredmeans of providing good contact between the rising gas and the downwardflowing liquid. Other standard type contacting means, such as sievetrays, packing, etc., may be used. More particularly, in the heattransfer section, it is advantageous to have a contacting means thatwill provide a relatively high contact area per vertical length of theabsorber. Thus packing, such as Glitsch Grid, may advantageously be usedin the heat transfer section. Water containing washed out absorbent anda small amount of soluble CO is Withdrawn at the bottom of the waterwash section through line 6 and passed to the lower part of stripper 41.Purified hydrogen is withdrawn from the top of the absorber in line 10at about 70 F.

Because the purified H withdrawn from the top of the absorber stillcontains some CO (between about 0.5-3.0 percent), one preferredembodiment of the present invention is to withdraw the H ,CO gas atbetween 15 and 45 F., generally about 35 F., and pass the H CO gas to aCuprous Ammonium Acetate (CAA) CO absorption system operating underrefrigeration. An added advantage of withdrawing the H CO gas at about35 F. is reduced methanol vaporization, thus allowing the water washsection to be eliminated or reduced.

CO -rich absorbent, for example methanol, is withdrawn from thebottom'of the absorber in line 12 and passed through pressure reductionvalve 13 and into the first letdown drum. The pressure on the richmethanol absorbent is reduced from about 1,500 p.s.i.a. to between 400and 1,200 p.s.i.a., generally to about 500 p.s.i.a. For a M s.c.f.d. Hproduction rate, the H and CO stream which flashes 01f from the methanoland is withdrawn in line 16 is about 4.6M s.c.f.d., consisting of about63 vol- 5 ume percent H 27 percent CO and the balance lighthydrocarbons, CO, argon, N etc. The gases withdrawn in line 16 arerecycled to join the feed gas in line 1.

Methanol withdrawn in line 18 from the bottom of the first letdown drumat about 75 F. is cooled by refrigerant in heat exchanger E-l. Therefrigeration duty is relatively small (only a few million B.t.u.s perhour for a hydrogen production rate of one hundred million standardcubic feet per day). Methanol withdrawn from E-1 in line 20 is furthercooled in E-2 and E-3 and then is flashed across pressure reductionvalve 25 into the second letdown drum at a pressure between about 3 and25 p.s.i.a. Temperature in the second let-down drum is about minus 90 F.As an alternative to pressure reduction valves 13 and 25, expansionturbines could be used to recover power.

CO withdrawn in line 26 from the top of the letdown drum is used toremove heat from the rich methanol absorbent passed through the heatexchangers B-1 and E-2 via line 20. The CO is withdrawn in line 28 atabout 60 F. Although refrigerant is shown as passing through a separateexchanger E-1 in the schematic flow diagram, the refrigerant mayadvantageously be passed through E-2 thus combining E1 and E-2 into asingle consolidated heat exchange unit. Since it is desired to flashlarge amounts of CO out of the absorbent in the second letdown drum andbecause CO is very soluble in methanol at temperatures 'as low as minus90 F., the partial pressure of CO above the cold methanol must bereduced below about 15 p.s.i.a.

The cold, lean absorbent is withdrawn from the bottom of the secondletdown drum in line 30 and passed in part through pump Pl to be fed vialine 3 to the absorber. The balance of the absorbent is passed via P2through exchanger E-3 into CO absorbent separator 36. Because the coldabsorbent passed via P2 through exchanger E-3 is heated in exchangerE-3, CO will be liberated from the absorbent upon entering separator 36.This cold CO at about F. is removed from the top of the separator 36 inline 38 and is passed through E-2 to join eflluent CO in line 28.

. Absorbent from the bottom of separator 36 is passed through exchangerE-2 and is thereby heated to about 70 F. It is then exchanged againstwarm, lean methanol at about 180 F. in line 40 by countercurrent passagethrough exchanger E-4. The methanol from separator 36 which has beenheated in E2 and 13-4 is passed at about 170 F. in line 42 to stripper10. The stripper 10 which is reboiled by steam reboiler E-6, elfectsfurther stripping of CO out of the methanol due to relatively hightemperatures, about 200 F., in the bottom of the stripper and due to therelatively low pressures, about 20 p.s.i.a. at the top of the stripper.The overhead of the stripper is comprised of a standard reflux systemwherein the overhead vapors from the column are partially condensed inheat exchanger E and the cooled reflux is returned to the column fromthe bottom of the reflux drum 44 via pump P-4 and line 46. CO iswithdrawn from the top of the reflux via line 48.

Stripped methanol is withdrawn from stripper via line 40 andrecirculated to the absorber via pump P-5. Water which has washed outabsorbent rising upward in absorber 2 and which has absorbed some CO iswithdrawn from the absorber in line 6 and fed to the lower part of thestripper. Because of the relatively high temperatures in the bottom ofthe stripper, the absorbent vaporizes and passes up in the stripper 10to be condensed in the cooler, upper part of the stripper and Withdrawnin line 40. The more volatile CO is not condensed in the upper part ofthe stripper but instead passes entirely through the stripper and iswithdrawn in line 48 as previously indicated. Water which is the leastvolatile component in the stripper, is withdrawn in a purified state,that is, substantially free of methanol and CO in line 50. The water isrecirculated to absorber 2 via P-6 after 6 it has been substantiallycooled by cooling water in exchanger E-7.

Referring now to 'FIGURE 2, a feed gas comprised of hydrogen, CO and H 5obtained, for example, by partial oxidation of a sulfur-containinghydrocarbon followed by CO shift conversion to hydrogen and CO is fedvia line 1, at a pressure between 900 p.s.i.a. and 10,000 p.s.i.a., forexample 1500 p.s.i.a., into the bottom of the absorber 2. The absorptionsection, heat transfer section, and water wash section of absorber 2 inFIGURE 2 operate similarly to those of the absorber in FIGURE 1 with thenotable exception that in FIGURE 2, H 8 is removed in addition to CO inthe absorption section. Purified hydrogen is withdrawn from the top ofthe absorber in line 10 and passed, for example, to a high pressurehydroconversion process, generally after removal or methanation of COcontained in the purified hydrogen withdrawn in line 10.

In the absorption section of absorber 2, H 8 as well as CO are absorbedand condensed at high pressure and then withdrawn from the bottom of theabsorber via line 12 into the first letdown drum. Hydrogen plus CO whichis flashed from the rich methanol fed to the first letdown drum iswithdrawn in line 16 from the top of the first letdown drum and recycledback with compression to join feed gas in line 1.

The absorbent withdrawn in line 18 from the bottom of the first letdowndrum is cooled in heat exchangers E-ll, E-12 and E-13 and then passedthrough pressure reduction valve 25 into the second letdown drum atabout 3-25 p.s.i.a. Cold, lean absorbent is withdrawn via line 30 fromthe bottom of the second letdown drum at about minus F. and passed vialine 3 into absorber 2.

A gaseous CO H 5 mixture flashed out of the methanol is removed from thetop of the second letdown drum via line 26 and passed into H 8 scrubber60". H 8 is scrubbed out of the gaseous CO H S mixture, and relativelypure CO is removed from the top of the H 8 scrubber via line 62 at about10 F. After heat exchange in E-12, which may be combined with E-11, theCO is withdrawn in line 64 at about 60 F.

The remaining part of the cold, lean methanol withdrawn in line 30 fromthe bottom of the second letdown drum is pumped via P2 through exchangerE-13 and into H S scrubber 60 via line 66. Because the H 8 is much moresoluble in the methanol absorbent than is CO there is a considerablygreater amount of H 8 than CO in the methanol absorbent passed via line66 to the H 8 scrubber. The greater solubility of H S in methanol andalso other hydrogen containing solvents, such as acetone, is probablydue to loose hydrogen bonding. Thus H S can be expected to be moresoluble than CO in any of the polar solvents containing hydrogen. Also,in addition to H 8 being more soluble in the absorbent, because the H Sis less volatile than is CO H S concentrates in the bottom of the H Sscrubber at about minus 10 F., whereas the more volatile CO leaves thetop of the H 8 scrubber at about 10 F. Reboiler heat to the H 8 scrubberis provided by heat exchanger E-15. H 3- rich methanol is withdrawn inline 68 from the bottom of the H 8 scrubber and is heated bycountercurrent heat exchange in heat exchangers E-14 and E-17. Operatingprocess-wise in this manner, relatively pure CO is obtained and thenecessity of using an inert gas, such as nitrogen, to strip CO from theabsorbent, is avoided.

Part of the H S-rich methanol from the bottom of H 8 scrubber 60 atabout minus 10 F. is fed to absorbent rectifier 70 via line 67 enteringthe rectifier near the top. The remaining part of the H S-Iich methanolfrom the bottom of H 5 scrubber 60 is heated in heat exchangers E-14 andE-17 and fed to rectifier 70 below the rectified absorbent withdrawalline 74. An H S-rich gaseous stream containing at least 10 percent H 8,and typically 15 to 30 volume percent H 8 is withdrawn in line 72 fromthe top of the rectifier. The stream is suitable for feed to a Clausplant for production of sulfur.

As the methanol flows downward in absorbent rectifier 70, it is heateddue to upward rising vapors generated by reboiler E18, thus stripping H8 and residual CO out of the absorbent. Rectified absorbent is withdrawnin line 74 from an intermediate point of absorbent rectifier '70. Therectified absorbent is substantially free of both CO and H S. The upperpart of rectifier 70 is provided with heat removal means, such asrefrigerant cooled condenser 80. One preferable mode of operation is tothus remove heat from the upper part of rectifier 70 and feed all of theH S-rich methanol to rectifier 70 at a point below the rectifiedabsorbent withdrawal.

After being cooled in heat exchanger E-17 by H rich methanol passingcountercurrently through the exchanger as indicated by line 68 and alsocooled by countercurrent cooling water in heat exchanger E-16, a part ofthe rectified absorbent at about 80 F. is passed via line 4 to anabsorber 2. The remaining part of the absorbent is further cooled inheat exchangers E-14 and E-15 and passed via line 76 to the upper partof the H S scrubber 60. This lean methanol absorbent, fed at about minus30 F. via line 76 to the top of the H 8 scrubber, scrubs H 8 out ofupward rising gases in the H S scrubber.

The upper part of the H 5 scrubber 60 is maintained at a pressure ofabout 16 p.s.i.a. The upper part of absorbent rectifier 70 is maintainedat about 45 p.s.i.a. Water, which is the least volatile component in themethanol- H S-WSIEI feed to absorbent rectifier 70 via line 69, isWithdrawn from the bottom of the absorbent rectifier via line 78. Thewater withdrawn in line 78, after cooling, is used as water wash forabsorber 2 overhead, in a circuit similar to that described with respectto FIGURE 1.

It is apparent that the invention has broad application to the removalof CO and other gases from gases which are not as highly soluble inpolar solvents. Also, those skilled in the art will appreciate that theterms cold and warm are relative and thus the particular temperaturesgiven in the detailed description wherein methanol was the exampleabsorbent may vary for other solvents while still being relatively coldand Warm for the particular solvent system set of temperatures. Althoughvarious specific embodiments of the invention have been described andshown, it is to be understood that they are meant to be illustrativeonly and not limiting. Certain features may be changed without departingfrom the spirit or essence of the invention. Accordingly, the inventionis not to be construed as limited to the specific embodimentsillustrated but only as defined in the following claims.

I claim:

1. In a process for removing CO from a CO -rich gas stream containing atleast one component having a boiling point lower than CO by absorbingthe CO in the CO -ricl1 gas stream fed to a lower point in an absorberusing an absorbent at low temperatures and high pressures, theimprovement which comprises:

(1) feeding cold, Co -lean, liquid absorbent to an intermediate point inthe absorber;

(2) feeding warm, lean, liquid absorbent to an upper point in theabsorber; and

(3) directly contacting cold, CO -Iean, gas, flowing upward in theabsorber past said intermediate point, with warm, Co -lean, liquidabsorbent flowing downward in the absorber so as to remove heat from theabsorbent and warm the cold, Co -lean gas.

2. A process for removing CO from H which comprises:

(1) absorbing CO in a cold polar absorbent so as to obtain cold CO -leanH and warm CO -rich absorbent;

(2) reducing the pressure on the warm CO -rich absorbent to obtain coldgaseous CO and cold CO lean liquid absorbent;

(3) heating and stripping further CO from the cold CO -lean liquidabsorbent to obtain warm Co -lean liquid absorbent;

(4) feeding the warm CO -lean liquid absorbent to an upper section ofthe absorber; and

(5) directly contacting the cold CO -Iean H rising upward in theabsorber with the warm CO -lean liquid absorbent falling downward in theabsorber to obtain CO -lean H at a temperature of at least 15 F.

3. Process as in claim 2 wherein the absorbent is propylene carbonate.

4. Process as in claim 2 wherein the absorbent is acetone.

5. Process as in claim 2 wherein the absorbent is methanol.

6. Process according to claim 5 wherein the feed gas is an H containingfeed gas at a pressure between 900 p.s.i.a. and 10,000 p.s.i.a.

7. Process according to claim 6 wherein the temperature of the feed gasis between 70 F. and 130 F. the temperature of the warm CO -richmethanol is between 50 F. and 110 F., the temperature of the cold CO-lean liquid methanol is between 30 F. and 1l0 F., and the temperatureof the warm CO -lean methanol is between 50 F. and 110 F.

8. In a process for removing CO and H 8 from a C0 and H S-rich gasstream containing at least one component having a boiling point lowerthan CO and H 5, by absorbing the CO and H 5 in the CO and H S-rich gasstream fed to a lower point in an absorber using an absorbent at lowtemperatures and high pressures, the improvement which comprises:

(1) feeding cold lean liquid absorbent to an intermediate point in theabsorber;

(2) feeding Warm lean liquid absorbent to an upper point in theabsorber; and

(3) directly contacting cold CO and H S-lean gas,

flowing upward in the absorber past said intermediate point, with warmlean liquid absorbent flowing downward in the absorber so as to removeheat from the absorbent and warm the cold CO and H S-lean gas to atemperature of at least 15 F.

9. Process according to claim 8 wherein the feed gas is an H containinggas at a pressure between 900 p.s.i.a and 10,000 p.s.i.a.

10. A process for removing CO and H 5 from an H containing feed gas*WhlCh comprises:

(1) absorbing the CO and H 8 in an absorber with a cold CO H S-Ieanpolar absorbent obtained as described in step 2, so as to obtain a coldCO H 8- lean, H -rich gas stream and warm CO H S-rich absorbent;

(2) reducing the pressure on the warm CO H S-rich absorbent so as toobtain a cold gaseous CO and H 8 stream, and the cold CO H S-Ieanabsorbent used to absorb CO and H 5;

(3) scrubbing H 5 from the cold gaseous CO and H 8 stream using leanabsorbent obtained as described in step 4, so as to obtain a Co -richgaseous stream and H S-rich absorbent;

(4) rectifying the H S-rich absorbent to obtain the lean absorbent usedto scrub H 5 and a gaseous H S- rich stream and a water stream;

(5 directly contacting the cold CO H S-lean, H -rich gas stream risingupward in the absorber with the warm CO H S-Iean absorbent streamflowing downward in the absorber.

11. Process as in claim 10 wherein the absorbent is methanol. 12.Process according to claim 10 wherein the feed gas is at a pressurebetween 900 p.s.i.a. and 10,000 p.s.i.a.

13. Process according to claim 12 wherein the temperature of the feedgas is between 70 F. and F., the temperature of the warm CO H S-richmethanol is between 50 F. and F., the temperature of the cold CO HS-lean methanol is between -30 F. and ll0 9 10 F., and the temperatureof the warm CO H s-lean 2,863,527 12/1958 Herbert et a1 6217X 2,926,7513/1960 Kohl et a1 5568 X methanol is between 50 F. and 110 F.

Refer nces cite REUBEN FRIEDMAN, Primary Examiner.

UNITED STATES PATENTS 5 R. W. BURKS, Assistant Examiner.

2,756,841 7/1956 Asendorph 5568X US, Cl, X R 2,781,862 2/1957 Fussman5568 X 5548, 73, 94

